I. Field of the Invention
The present invention relates to an apparatus and reactor system for recovering polymer grade olefins. Particularly, the invention uses a unique process ordering to remove contaminants that would otherwise prevent the recovery unit from producing polymer grade ethylene and propylene.
II. Background of the Related Art
The mono-olefinic compounds, such as ethylene and propylene, are important commodity petrochemicals useful in a variety of processes for making plastics and other chemical compounds. For instance, ethylene is used to make various polyethylene plastics and chemicals such as vinyl chloride, ethylene oxide, ethylbenzene and alcohol. Propylene is used to make various polypropylene plastics and chemicals such as acrylonitrile and propylene oxide.
Traditionally, the olefins have been produced from petroleum feedstocks at high temperature by steam-cracking, fluid catalytic cracking or deep catalytic cracking processes. See, for example, Hallee et al., U.S. Pat. No. 3,407,789; Woebcke, U.S. Pat. No. 3,820,955, DiNicolantonio, U.S. Pat. No. 4,499,055; Gartside et al., U.S. Pat. No. 4,814,067; Cormier, Jr. et al., U.S. Pat. No. 4,828,679; Rabo et al., U.S. Pat. No. 3,647,682; Rosinski et al., U.S. Pat. No. 3,758,403; Gartside et al., U.S. Pat. No. 4,814,067; Li et al., U.S. Pat. No. 4,980,053; and Yongqing et al., U.S. Pat. No. 5,326,465.
It is also well known in the art that mono-olefinic compounds can be produced from oxygenates, especially alcohols. There are numerous technologies available for producing oxygenates including fermentation or reaction of synthetic gas derived from natural gas, petroleum liquids, carbonaceous materials including coal, recycled plastics, municipal waste or any other organic material. Generally, the production of synthesis gas involves a combustion reaction of natural gas, mostly methane, and an oxygen source into hydrogen, carbon monoxide and/or carbon dioxide. Syngas production processes are well known and include conventional steam reforming, autothermal reforming or a combination thereof. Methanol, the preferred alcohol for light olefin production, is typically synthesized from the catalytic reaction of hydrogen, carbon monoxide and/or carbon dioxide in a methanol reactor in the presence of a heterogeneous catalyst. For example, in one synthesis process, methanol is produced using a copper/zinc oxide catalyst in a water-cooled tubular methanol reactor. The conventional methanol conversion process is generally referred to as a methanol-to-olefin(s) (MTO) process, where methanol is converted to primarily ethylene and/or propylene in the presence of a molecular sieve catalyst.
Unfortunately, in addition to the mono-olefins obtained by these processes, the produced gases typically contain a large amount of other components such as diolefins, hydrogen, carbon monoxide and paraffins. While, the content of these compounds depends on the severity of the conversion treatment, it is often too low to economically justify their separation and use. Nonetheless, even in low quantities some of these byproducts/contaminants, such as, for example, acetylene and methyl acetylene, must be removed because they act as a poison to the catalysts used for making polymers, e.g., polyethylene, polypropylene, out of the mono-olefinic compounds. Another trace contaminant that may cause a problem in polymer production are oxygenated hydrocarbons (ethers, esters, acids, carbonyls) because the polar nature of the oxygenated hydrocarbons will deactivate the Ziegler-Natta or metallocene polymerization catalysts.
To this end, treating the olefin stream in order to obtain polymer grade ethylene/propylene can take various routes depending on the level of an acetylenic compounds and oxygen present. For instance, plural stage rectification and cryogenic chilling trains have been disclosed, for example in Perry's Chemical Engineering Handbook (5th Edition) and other articles on distillation techniques. Typical rectification units are described in Roberts, U.S. Pat. No. 2,582,068; Rowles et al., U.S. Pat. No. 4,002,042, Rowles et al., U.S. Pat. No. 4,270,940, Rowles et al., U.S. Pat. No. 4,519,825; Rowles et al., U.S. Pat. No. 4,732,598; and Gazzi, U.S. Pat. No. 4,657,571. Especially successful cryogenic operations are disclosed in McCue, Jr. et al., U.S. Pat. No. 4,900,347; McCue, Jr., U.S. Pat. No. 5,035,732; and McCue et al., U.S. Pat. No. 5,414,170.
In a typical hydrogenation reactor, the content of acetylenes in the olefin stream is reduced to below 10 ppm, wherein the acetylenes are converted via (i) catalyst adsorption; (ii) isomerization and (iii) hydrogenation steps to ethane, propane, ethylene or propylene. The process employs hydrogen, which contacts one or more acetylenes under conditions effective to selectively hydrogenate the acetylenes, thereby forming a substantially acetylene-free stream. Preferably, after the hydrogenation of acetylenes, the catalyst does not further hydrogenate the partially hydrogenated product, i.e., ethylene or propylene. For example, the hydrogenation of ethylene and propylene must proceed much more slowly than the partial hydrogenation of acetylenes. Moreover, the partially hydrogenated product can also be protected from subsequent reactions by being rapidly desorbed from the catalyst surface and then not being re-adsorbed again. This thermodynamically dependent selectivity is based on the triple bond being more strongly adsorbed than the corresponding double bond, because of its more electrophilic character. Even relatively small differences in the adsorption energy are sufficient for an acetylenic compound to immediately displace the primary resulting hydrogenation product from the catalyst surface and accordingly act as a retardant for the subsequent reactions.
Conventional hydrogenation processes for the selective hydrogenation of acetylene in the presence of ethylene are typically supported by metals of Group VIII, of which palladium was shown to be the most active and selective metal for the hydrogenation of acetylenes to the corresponding olefins. By-products of the olefin production plant, such as hydrogen sulfide (H2S) and high concentrations of carbon monoxide (CO), typically are removed prior to palladium catalyzed hydrogenation, as these by-products can poison the palladium catalyst. In process configurations where H2S and CO are still present, use of composites of Group VIII (Co or Ni) and Group VIB (Mo or W) metals in a sulfided state has been proposed, one of which is a cobalt or nickel molybdenum. For example, see U.S. Pat. No. 3,649,525, incorporated herein by reference, wherein a Ni/Mo catalyst was used for desulfurization and hydrogenation of heavy oils. For other examples, numerous references are provided in Sulphide Catalysts: Their Properties & Applications by Weisser and Landa (Pergamon Press, 1974), which is incorporated herein by reference. Alternatively, the use of the copper catalysts has also been proposed because it is highly selective to acetylenic hydrogenation. Unfortunately, the activity of copper catalysts is too slow and the catalyst cycle time is undesirably short for the feed streams, which contain higher than about 2000 ppm total alkynes due to fast deactivation caused by the deposition of polymeric material on the catalyst surface.
In a typical conventional separation process, as shown in FIG. 1A, the source such as cracked gas or OTO (oxygenate to olefin) reaction product in a line 101 is compressed in a compressor 201. The compressed gas in line 102 is then caustic washed in the AGR reactor/washer 202 and fed via a line 103 to dryer 203. The dried gas in line 104 is then fed to the chilling unit 205. The liquids from the chilling unit 205 are removed via a line 106.
In the back-end configuration (referring to the placement of the hydrogenation tower), as shown in FIG. 1B, the liquids from line 106 are fed to a demethanizer tower 206. The methane and lighter components are removed from the top of the demethanizer tower 206 via line 108 for further processing. The C2+ components are removed from the bottom of the demethanizer tower 206 via line 107 and fed to a deethanizer tower 207. The C2 components are removed from the top of the deethanizer tower 207 in a line 109 and passed to an acetylene hydrogenation reactor 208 for selective hydrogenation of acetylenes and subsequently to a caustic tower 209 to remove H2S if the sulfide catalyst has been used. The effluent from the hydrogenation reactor 208/optional caustic tower 209 is then fed directly via a line 113 to a C2 splitter 210 for separation of the ethylene from ethane. The ethylene is removed from the top of splitter 210 via C2 heat pump 211 in a line 114 and ethane is removed from the bottom of splitter 210 in a line 115. The C3+ components removed from the bottom of the deethanizer tower 207 in a line 110 are directed to a depropanizer tower 212. The C3 components are removed from the top of the depropanizer tower in a line 111 and fed to a C3 hydrogenation reactor 215 to selectively hydrogenate the methyl acetylene and propadiene. If the sulfide catalyst has been used, the effluent from the hydrogenation reactor 215 is passed through a caustic tower 216 and then fed via a line 112 to a C3 splitter 213, wherein the propylene and propane are separated. The propylene is removed from the top of the C3 splitter in a line 116 and the propane is removed from the bottom of the C3 splitter in a line 117. Finally, the C4+ components removed from the bottom of the depropanizer tower 212 in a line 118 are directed to a debutanizer 214 for separation into C4 components and C5+ gasoline. The C4 components are removed from the top of the debutanizer 214 in a line 119 and the C5+ gasoline is removed from the bottom of the debutanizer 214 in a line 120.
In the front-end configuration, as shown in FIG. 1C, the liquids from line 106 are fed to a depropanizer tower 212. The C3− components are removed from the top of the depropanizer tower in a line 111 and fed to a hydrogenation reactor 208 to selectively hydrogenate the acetylene, methyl acetylene and propadiene. If the sulfide catalyst has been used, the effluent from the hydrogenation reactor 208 is passed through a caustic tower 209 and then fed via a line 113 to the demethanizer tower 206. The methane is removed from the top of the demethanizer tower 206 via line 108 for further processing. The C2 and C3 components are removed from the bottom of the demethanizer tower 206 via line 107 and fed to a deethanizer tower 207. The C2 components are removed from the top of the deethanizer tower 207 in a line 109 and passed to a C2 splitter 210 for separation of the ethylene. The ethylene is removed from the top of splitter 210 via C2 heat pump 211 in a line 114, and ethane, removed from the bottom of splitter 210 in a line 115. The C3 components removed from the bottom of the deethanizer tower 207 in a line 110 are directed to a C3 splitter 213, wherein the propylene and propane are separated. The propylene is removed from the top of the C3 splitter in a line 116 and the propane is removed from the bottom of the C3 splitter in a line 117. Finally, the C4+ components removed from the bottom of the depropanizer tower 212 in a line 118 are directed to a debutanizer 214 for separation into C4 components and C5+ gasoline. The C4 components are removed from the top of the debutanizer 214 in a line 119 and the C5+ gasoline is removed from the bottom of the debutanizer 214 in a line 120.
In another permutation of the front-end configuration, the depropanizer tower 212 is placed after the hydrogenation step. Specifically, the liquids from line 106 are fed initially to a hydrogenation reactor 208 to selectively hydrogenate the acetylene, methyl acetylene and propadiene. The effluent from the hydrogenation reactor 208 is then passed through a caustic tower 209 and fed via a line 113 to the rest of the processing/distillation units. This approach is commonly referred to as “raw gas configuration.” Commonly, a full-range stream containing both light and heavy components ranging from hydrogen up to C5's and heavier is processed over a fixed bed of selective hydrogenation catalyst, preferably the partially sulfided nickel catalyst. This catalyst is operated to effect complete removal of simple acetylene and removal of a majority of the methyl acetylene and propadiene. However, about one-half of the butadiene is also hydrogenated. Because these catalysts also promote the hydrogenation of butadiene, they cannot be used when butadiene is in the system unless the C4 stream is first removed. In such instances a depropanizer is generally employed before subjecting the stream to a hydrogenation reaction (front-end process; see FIG. 1C). Furthermore, the use of a raw gas catalytic hydrogenation reactor has not been wide practiced because the position in the flowsheet requires the reactor to fully hydrogenate acetylene, methyl acetylene and propadiene, which all have different rate kinetics. Therefore, subsequent hydrogenation is still required in order to meet product specifications, e.g., the production of the polymer-grade olefin. The present invention solves this impracticality because the raw gas stream in an oxygenate to olefins or MTO process has significantly lower concentrations of acetylene, methyl acetylene and propadiene and makes the raw gas hydrogenation process a viable commercial option.
The front-end reactor configuration of the prior art, however, has its own shortcomings. The process requires multiple beds to reduce temperature rise because the plant experiences operating upsets due to temperature excursions during the initial start up resulting from the sensitivity and activity of the fresh catalyst. Furthermore, the hydrogen to acetylene ratio is uncontrollable. On the other hand, in the “back-end” reactor configuration of the prior art, the process requires an external fluid solvent and must carefully regulate the hydrogen ratio, carbon monoxide content and reactor inlet temperature due to pressure sensitivities to excursions in acetylene and carbon monoxide concentrations. The reactor effluent typically contains less than 1 ppm of acetylene but is contaminated with traces of hydrogen and methane, which also represents a major disadvantage.
Essentially, none of the prior art processes have described a useful method of obtaining relatively high purity olefin components, i.e., polymer grade ethylene and propylene (˜99.9% purity), from olefin-containing streams such as an effluent from various types of fluid catalytic cracking reactors in a raw gas catalytic hydrogenation reactor and the effluent from the oxygenate-to-olefin reactors such as the MTO. Therefore, it is desirable to have a useful method of obtaining relatively high purity olefin components from olefin-containing streams that do not have the drawbacks of the systems of the prior art, such as in the use of front-end or back-end hydrogenation systems.